Process for isomerizing normal paraffins



Sept. 6, 1960 H. o. FOLKINS ETAL 2,951,887

PROCESS FOR ISQMERIZING NORMAL PARAFFINS Filed April 14. 1958' 2 Sheets-Sheet 1 8 "R0 V E E at mm w w m T 55$: ommy A 22 s fi WW mum w v I' x Y no: Do- B 2 y 8 Q G v m5 2 2 mm m. V x mm 3 U m 5 7 Q a m uu. H 3 T n H V X Z N V 3 u. m m a 1. u m v in 0N m 2 on y no tm mo zmoomtz mi United States Patent l rnocnss FOR ISOMERIZING NORMAL PARAFFINS Filed Apr. 14, 1958, Ser. No. 728,467

'6 Claims. (Cl. 260--683.65)

This invention relate-s to the catalytic hydroisomerization of isomerizable hydrocarbons having 4 to 7 carbon atoms per molecule. It is more specifically concerned with improving the octane rating of petroleum hydrocar- 'bon feed stocks consisting predominantly of normal hexane and/or normal pentane hydrocarbons.

According to this invention, it has been found that the hydroisomerization of hydrocarbon feed stocks consisting predominantly of isomerizable aliphatic and alicyclic hydrocarbons having 4 to 7 carbon atoms per molecule can be efficiently carried out by processing the feed stocks at a temperature within the range of about 600 to 775 F., a pressure within the range of 100 to 1000 p.s.i.g., and a hydrogen/hydrocarbon mol ratio within the range of about 0.5 to 5 in the presence of a composite catalyst consisting essentially of a rhodium promoted silica-alumina composite.

In integrated petroleum refining operations for the production of high-octane-nnmber gasolines, in order to obtain maximum eilectiveness, one of the unit processes selected must be for the processing of feed stocks consisting predominantly of the loWer-molecular-weight, nor 'mally liquid, saturated aliphatic and alicyclic hydrocar- T'bons containing 4 to 7 carbon atoms per molecule. Sub- ;stantial quantities of these feed stocks are available to "warrant the separate processing of these materials. Al-

though octane number improvement can be obtained by treating these feed stocks in a dehydrogenation process to produce olefins, it is more desirable from an octane-yield relationship to utilize isomerization processes for effecting the octane number improvement in these compositions. Furthermore, the isomerization product has an increased research octane number and improved -road performance, and is a stable product which augrnents the stability of the blended, finished gasoline.

Because of the importance of isomerization as a unit process in an integrated refining scheme for the production of high-octane-number gasolines, several commercial isomerization processes have been developed which utilize solid catalysts. The use of such catalyst eliminates plant corrosion problems and the accompanying thigh maintance cost which are attendant upon the use of catalysts of the Friedel-Crafts type. Because of the effectiveness of platinum-promoted catalyst composites such as platinum-h-alogen-alumina, platinum-silioa-alumina, etc., in hydroforming'operations, it has been natural to utilize similar catalysts in isomerization processes These catalysts, however, require high operating temperatures which .are disadvantageous because isomeriza- 'tion is an equilib-rum reaction, the efiiciency of which decreases with an increase in the processing temperature. As a result, at the high temperatures employed, the equilibrium product contains substantial quantities of lowoctane-number paraflins which are notisomerized during the course of the reaction. In addition, it has been reported that as the equilibrium conversion is ap preached, the gas loss increases very sharply. Further 2,951,887 Patented Sept. 6, 1960 disadvantages resulting from the use of high-temperature processing conditions are higher fuel cost for carrying 'out the reaction, and added expense for fabricating process vessels to withstand the combination of high pressures and high temperatures required for this type of isomerization. A non-platinum-containing, noble-metal catalyst, however, has been found which permits the isomerization reaction to be carried out at lower operating temperatures and thus avoids the disadvantages accompanying high-temperature isomerization using platinum catalysts.

It is, therefore, the primary objective of this invention to provide a hydroisomerization process for the isomerization of hydrocarbon feed stocks consisting predominantly of loW-molecular-weight isomerizable saturated hydrocarbons having 4 to 7 carbon atoms per molecule at temperatures not higher than about 775 F. It is another object of this invention to provide a low-tem perature hydroisomerization process for improving the octane number of petroleum-derived feed stocks consisting essentially of C -C normal paraffinic hydrocarbons. It is an additional object of this invention to carry out the hydroisomerization process employing a solid, non-corrosivecatalyst utilizing relatively low temperatures, not higher than about 775 R, which permit the substantial production of branched-chain isomerization products with a minimum loss to gaseous products con sisting of butanes and lower-molecular-weight hydrocarbons.

tion of saturated, aliphatic and alicyclicyisomerizable hydrocarbons having 4 to 7' carbon atoms per molecule. These catalysts operate efficiently at reaction tempera- In carrying out the process of this invention, a composite catalyst containing 0.0I1.0% by weight of rhodium and preferably between about 0.10.4% by weight, based on total catalyst composition, incorporated in a silica-alumina hydrocarbon cracking catalyst containing more than 50% by weight of silica, based on said cracking catalyst, is employed to carry out the hydroisomerization of saturated isonierizable hydrocarbons having 4 to 7 carbon atoms per molecule, employing the following operating conditions:

Range Preferred Range em erature F. for

Pressure, p.s -1, 000 350-750 Liquid Hourly Volume Space Velocity O. 5-10 1-4 lib/hydrocarbon n10} ratio 0. 5-5 1. 5-4. 5

1 The liquid volume, at 60 F., of limiting reactant fed per hour per unit volume of efiective catalyst bed.

In the isomerization process of this invention, a variety of processing schemes are available. In Figure 1 is shown a simple scheme which utilizes a feed preparation and product recovery system employing a minimum number of process towers. A light, straight-run naphtha feed having as ASTM boiling range of about 100-180 F. is introduced into a deisohexanizer 10 via line 11.

, The residue consisting essentially of n-h-exane and heavier 3 hydrocarbons is sent through line 12 to reactor 13 for isomerization. The isohexa'fie and lighter hydrocarbons are removed from deisohexanizer and transferred by means of line 14 to depentanizer 15 where the isohexane and heavier hydrocarbons are separated and removed from the system via line 16 to storage. The overhead from depentanizer 15, which consists essentially of normal and isopentane, is sent to C -splitter 18 through line 19. Isopentane is recovered in the fractionator overhead, and transferred to storage through line 20 or to suitable gasoline blending facilities (not shown), and the residue consisting predominantly of normal pentane is transferred by means of line 21 to a point of confluence with line12 wherein it is sent to reactor 13 for processing. The reaction effluent is transferred through line 23 to stabilizer 22 where the butane-and-lighter fraction is separated. The pentane-and-heavier fraction is then processed in deisohexanizer as described above.

Catalysts employed in carrying out the process of this invention are prepared by conventional techniques by incorporating a small amount of rhodium in a hydrocarbon cracking catalyst consisting of a silica-alumina composite containing more than 50% silica. The small amount of rhodium is incorporated in the silica-alumina support by impregnation of the support with a reducible soluble salt or complex of rhodium, e.g., an aqueous solution or an acid solution of the metal chlorides, such as a hydrochloric acid solution of rhodium chloride or nitrate, mixed salts, e.g., ammonium chlororhodinate, or nitrito, amino, and nitritoamine complexes of rhodium. Colloidal solutions of salts of rhodium can also be used as the impregnant. The preparation of the supported catalyst is generally carried out by wetting the support with an aqueous solution of rhodium chloride. The quantity of rhodium incorporated in the silica-alumina support can vary between about 0.0l-1% by weight, based on catalyst composition. It has been found that optimum efliectivenessris obtained if an amount within the range of 0.1-0.4 wt. percent is employed.

Prior to impregnating the silica-alumina support, it is preferred that this component of the catalyst composition be dried at an elevated temperature within the range of about 250-400 F. The impregnated mass is dried for at least about four hours at a temperature within the range of 225-35 0 F. The green catalyst is then pelleted and activated by contacting the dried catalyst mass with a stream of hydrogen at an elevated temperature for a time sufiicient to reduce the rhodium component of the green catalyst to the metallic state. This reduction can generally be effected by heating the catalyst mass to a temperature between 850-975 F. in hydrogen over a period of about 2-24 hours. In general, 2000 to 5000 s.c.f.h. of H per barrel of catalyst are used in this reduction step. Although the preparation of the supported rhodium catalysts employed in the process of this invention is generally carried out in accordance with the foregoing procedure, other manipulative procedures for incorporating the desired amount of metallic rhodium on the selected silica-alumina catalyst base can be used, including techniques wherein a combination of oxidation and reduction steps are used to provide the desired catalyst composition.

In selecting the silica-alumina hydrocarbon cracking catalysts for use as supports in the preparation of the isomerization catalyst employed in the process of this invention, it is necessary that the composition contain not less than about 50% by weight of silica. Accordingly, this component of the isomerization catalyst composition has a silica content within the range of about 50- 95% by Weight, preferably 75-90%, and an alumina content within the range of about 50-5%, preferably 25-10%, and is composited to evince acidic properties and hydrocarbon cracking'activity. The silica-alumina support can be obtained commercially or can beprepared by admixing separately prepared portions of Silica gel and alumina gel, or in the alternative, by conventional co-precipitation techniques. It is also possible to prepare a catalyst which can be employed in the instant invention by contacting silica gel particles with a solution of an aluminum salt and a rhodium salt of the desired concentrations. After drying the mixture, it is heated for a sufficient time toeft'ect the decomposition of the salts. Thereafter, the rhodium is reduced to the metallic state by treatment with hydrogen at elevated temperatures.

The process of this invention is especially'adaptable for effecting the isomerization of feed stocks consisting predominantly of normal pentane and/ or normal hexane to produce an octane improvement by promoting the molecular rearrangement of these hydrocarbons, or mixtures thereof, such as light petroleum fractions having an ASTM boiling range of -200 F.

The invention is further illustrated by the following non-limiting and illustrative example of the u'se'of a rhodium-promoted, silica-alumina, composite catalyst in the isomerization of low-boiling isomerizable hydrocarbons having 4 to 7 carbon atoms.

A rhodium-promoted, silica-alumina catalyst was pre pared by impregnating a silica-alumina support having the following composition:

Weight percent Composition:

A1203 Na 'O 0.021 S04 l Fe 0.25 SiO 75.3 1

with an acidified, aqueous solution of rhodium chloride. The rhodium chloride solution was prepared by dissolving 1.0 gram of rhodium chloride in distilled water. To prepare a catalyst containing 0.2% by Weight (based on total catalyst composition) of rhodium on a silica-alumina support consisting nominally of 75% by weight of silica and 25% by weight of alumina,250 grams of the silicaalumina were placed in a suitable vessel containing 250 ml. of rhodium chloride solution. This volume-of solution represents the average amount necessary to fill the actual pore volume of the silica-alumina support employed. If some other silica-alumina is employed, a volume of solution compatible with the adsorptive capacity of that support should be used. The impregnated support was removed from the flask, dried at about 250 F. for sixteen hours, and subsequently pelleted into A; x inch pellets. Thereafter, the dried mass was activated by heating to 975 F; in hydrogen over a period of five hours, followed by continued treatment with hydrogen at 975 F. for sixteen hours. After purging with nitrogen and cooling to 750 F., the catalyst was oxidized with air for one hour. The-catalyst, when cool, was then placed in the reactor and heated to 975 F. in hydrogen. Thereafter, the catalyst was treated with 6-8 s.c.f.h. of hydrogen at 975 F. for eight to sixteen'hours, after which it was cooled to 700 F. Following this, the reactor was pressurized to reaction pressure with hydrogen and the hydrocarbon feed stock was charged under desired conditions. For catalysts of best activity, it is preferred that the rhodium be added to a support which has not been subjected to temperatures above a certain limit at any time in the history of the support. Thus, the support should not have been heat-treated beyond an equilibrium dehydration temperature of about 800-900" F., and the water content should not have been reduced below about 1% of the catalytic support.

A tabular summary of data obtained employing the foregoing catalyst composition, as well as othercatalyst compositions similarly prepared, is set forth in Table I. Also included in Table I are comparative data demonstrating the superiority-of the instant rhodium-promoted, silica-alumina over similar catalysts employing platinum as the promoting agent. i

Table I Catalyst, wt. Percent:

Promoter Support Feed, Wt. Percent:

Pri ei-Co Run Conditions:

Pressure, p.s.i.g L H V H /Hc Mol. ratio 0.2 Pd 0.2 Pt

Temp, F Conversion:

1'05 yield, wt. percent.-

Selectivity Catalyst, wt. percent:

Promoter 1 Support Feed, wt. percent:

eun'etinsiaa'sz Pressure, p.s.i.g LWH SV HlHc Mol. ratio Temp, "F Conversion: i-C yield, wt. percent Selectivity r7595 emf-% A1103.

The advantages that are to be obtained from employing the process of this invention are clearly shown by the data presented in Table I. It is seen that, at lower concentrations of metal, the rhodium-promoted catalysts are far more effective than similar catalysts promoted with ruthenium, or palladium for the isomerization of n-paraflin-containing feed stocks. It is also evident that no significant change in yield occurs with increase in the amount of rhodium present in the catalyst.

To obtain maximum efficiency, auxiliary equipment is employed for pretreating the feed stock and the hydrogen utilized in the isomerization process. In order to insure long catalyst life, it is necessary to employ a hydrocarbon feed stock which is substantially free from sulfur or sulfur-containing compounds. Accordingly, when necessary, a pre-treater or guard-case should be installed in the feed line immediately ahead of the reactor to effect the removal of substantially all the sulfur compounds from the feed. The guard-case is at substantially the same total flow and operating conditions exceptfor pressure, employed in the isomerization. The desulfurization before the guard-case can be effected by acatalytic, vaporphase, desulfurization process in the presence of clay, bauxite, cobalt molybdate, nickel molybdate, nickel, or other suitable catalysts for effecting the desulfurization of the feed stock in the presenceof hydrogen. A variety of desulfurization methods based upon the decomposition of the sulfur compoundsat elevated temperatures and in the vapor phase are briefly described by Kalichevsky, Petroleum Refiner, vol. (4), at page 117, et seq. When the reactant-catalyst in a guard-case is saturated with sulfur, i.e., when sulfur appears in the efiluent, it is by-passed and prepared for regeneration or removal of sulfur by high-temperature oxidation in the presence of a steam carrier-gas. The process need not be interrupted or shutdown, provided the sulfur content of the feed is less than about 3 p.p.m. Two to ten parts per million is about the maximum level normally expected in the catalytically- -desulfurized feed to the guard-case unit. The regeneration of the reactant-catalyst in the guard-case is carried out in the same manner as for hydrodesulfurization catalysts. A small furnace is needed for preheating the steam. The final oxidation temperature should be about 1000 F. After gross sulfur removal is complete, steam and air are removed from the vessel by evacuation and/or purging. The unit is then returned to normal processing conditions. The operating temperature of the guard-case can be from about 550725 F. Employing tempera tures substantially higher than 725 F. causes excessive hydrocracking. The amount of reactant-catalyst employed in the guard-case depends upon the following factors: (l) the amount of sulfur in the feed to this unit, (2) the length of the processing cycle between regenerations of the reactant-catalyst, and (3) the total flow rate. The amount of sulfur that the guard will hold is calculated on the assumption that one atom of sulfur is held by one atom of metal. The efficiency factor to be used with this relation should be in the range of 70-90%.

It is also highly desirable, for maximum efficiency, that the hydrogen employed as a processing aid in the hydroisomerization' process be substantially free of H 0, 0 CO, H 8, and related compounds, including those which react under hydroisomerization conditions to form these materials. Although it is preferred that the hydrogen be free of these impurities, trace amounts of these substances not in excess of about 2 parts per million can be tolerated. The hydrogen employed can be obtained from conventional sources. Commercial isomerization systems, however, will for reasonsof convenience and efiiciency operate in conjunction with conventional hydroforming units. In these instances the hydrogen-rich gas from the hydroformer can be employed not only for makeup in the isomerization reaction section, but also for putting the unit on stream. Although the small quantities of dry hydrocarbons in this hydrogen-rich stream may result in a small loss in yield at the same space velocity, the yield can be held constant by an appropriate change in space velocity.

The operating conditions for the process are normally selected so that degeneration or fouling of the catalyst: does not occur. However, degeneration may developfrom extremely long process periods or operational upsets. Operational upsets involving reactor temperatures. in excess of the design level and/or the loss of hydrogen circulation can cause catalyst fouling. In these cases,.a. simple regeneration in place is accomplished by the pro No deactivation has been en- 8 be' purehydrogen or reformer oflf-gas. The water partial pressure during reduction should be below about 15 mm.

' and preferably as dry as practicable. Time of reduction by misoperation or by exhaustive processing, and certain catalyst-water-high-temperature interactions. The presence of trace water, in excess of that likely from saturated feed streams at roughly 90 F, has no effect during processing conditions. However, high partial pressures of water, high temperature, and extended times at certain conditions cause the catalyst to become temporarily deactivated. These conditions should not be employed in the regeneration cycles. In the event difficulties arise, proper definition of procedure is needed. A simple regeneration cycle of the type defined herein will completely revive a catalyst which has lost activity due to the above causes. To define typical instances tobe avoided, the following examples are given:

(1) If the gases are dry, which constitutes less than about 1 mm. mercury partial pressure of water, the catalyst can withstand temperatures as high as 975' F. for extended times and up to 1000 F. for times less than about one hour;

(2) If the gases in the reactors contain mm. mercury partial pressure of water, then the upper limit of temperature and time is 975 F. for about one hour; and

(3) A partial pressure of 25 mm. or greater at 975 F. or higher cannot be tolerated.

Another unusual operational situation will cause catalyst fouling and will result in the need for catalyst regeneration. 'If the reactor is depressurized after a processing cycle, it is advisable that a simple regeneration be conducted. Apparently, coke lay-down is unavoidable in this situation.

It is emphasized that these conditions which cause difficul ty are abnormal and they can be avoided by proper design and operation of the reaction section.

The three functional variables involved in the regeneration are as follows:

(1) Condition of dryness of the catalyst during all regenerative treatmenzs.This is controlled by the water partial pressure and temperature within the reactor. The water partial pressure Within the catalyst bed should be as low as practicable and certainly below 15 mm. mercury.

(2) Oxidation tempe1rature.-The safe upper level of temperature is 950 F. Below this value and above about 700 'F., the oxidation temperature is not critical, except that it should be high enough for substantial coke oxidation. A burning front temperature of about 800- 850 F. and a final treatment with the inlet gas and entire bed at about 850 F. are recommended. The time requirementis consistent with oxidation completion and manipulations. Normally, the coke laid down is only about 0.1 to 1% of the catalyst in comparison to a representative metal content of 0.4%. The metal will co-mpletely oxidize with an oxygen partial pressure at any convenient level, e.g., l5% oxygen at 25 p.s.i.g. total pressure. The inert gas can be nitrogen, or carbon monoxide-free flue gas.

. (3) Reduction temperature-The safe upper level of reduction temperature is 950 F. The recommended reduction temperature is about 850 F. This factor is not critical in the range 750-950 F. The reduction gas can is relatively unimportant, and a time of about one hour at these conditions is suggested. It is also suggested that in the initial phase of the reduction, when water is liberated from the reduction of the oxide the total pressure and the circulation gas rate'should be such that only a low partial pressure of water exists in 'the catalyst bed at all times.

Other techniques can be employed in the design of the regeneration system. For example, first, reverse flow can be used in the reactor for purging, or evacuation, and for the initial burn-off. Since it is desirable to eliminate sulfur compounds and water from the gases and the catalyst, the flow of dry oxidation gas should be once-- through, with no recycle. The oxidation time and ca-. pacity are so low that this should be the easiest and best approach. Second, if an inert gas is employed, it should be dried in order to avoid possible difiiculty from water and to afford a means oftdrying the catalyst by purging. Third, regeneration connectionsto and around thereactor should include a minimum of the piping and equipment normally used for the feed and product streams. If a recycle gas compressor is to be employed during the regeneration, processing tanks, exchangers and process piping should be by-passed.

A catalyst regeneration after a long processing cycle can be carried out as follows;

(1) After the feed system has been shut down and the heat outputs have been decreased, the gas-recycle circulation is continued until the feed lines and reactor are essentially purged of hydrocarbons at operating temperature and pressure. Gas circulation is then stopped.

(2) The complete unit is depressurized and the gas vented.

(3) After appropriate valving changes, the isolated reactor is dried or purged with dry, inert gas. Evacuation, as a method of drying and eliminating combustibles, can be substituted for the initial inert-gas purge. The catalyst-bed temperature is maintained around 750-800 F. for the purging step.

(4) After the purging step to remove combustible gas is complete, the oxidation step is commenced. When the bed and inlet gas temperatures are in the range of 750-' 800 F., air (oxygen) is introduced into the carrier gas to a content of about 1 to 5% oxygen. The mass rate of oxygen is controlled so that the oxidation-front temperatures do not exceed 850 F. When the initial burnoff is complete, the reactor temperature is lined out at about 850 F. in the presence of the oxygen-containing gas for a time consistent with thorough oxidation. At this time and thereafter, the presence of some moisture in the regeneration gases is not critical, except that water partial pressures should be held below about 15 mm.

(5) After the removal of oxygen from the reactor and system, the catalyst is reduced with hydrogen or hydrogen-rich gas at 850 F., at about one atmosphere and for a time period of about one hour. This concludes the regeneration cycle and the unit is pressurized, hydrogen circulation is established, and the bed temperature is adjusted to the defined level for processing conditions.

Referring to Figure 1, reactivation of the catalyst is provided for by means of the reactivation system shown schematically. Inert purging gas, such as nitrogen from line 30, is fed into line 31 through which air is admitted into the system. Before initiating regeneration, the reaction system is thoroughly purged free of hydrocarbons. Thereafter, the reaction system is depressurized to about atmospheric pressure, and the system is purged with heated nitrogen which is passed through .line 31 into reactor 13. The purge gases are removed from the system through vent 32. The purging can also be efiected by evacuating the system by means of steam ejector 33 which is connected to the system. through line 32. After .the

reactor system has been purged, controlled amounts of air are introduced into the flowing nitrogen stream by means of air supply-line 31 and the air-nitrogen mixture is passed through the reactor to oxidize the activated catalyst. This is followed by the reduction step abovedescribed. Then the reactor is cooled to the desired reaction temperature, the reaction system is repressured, and fresh feed is reintroduced into the reactor.

From the foregoing description of this invention, it is apparent that numerous combinations of reactors and fractionators are possible for carrying out the isomerization process of this invention for the processing of light hydrocarbon feed stocks. The process of this invention finds application in combination with other conventional unit refining processes, such as reforming, or in splitstream techniques employing a plurality of reactors to separately process feed stocks under isomerization conditions selected to obtain maximum efficiency with respect to the feed stock being processed. The various feed components can be processed jointly or singly, and on a oncethrough or recycle basis. In applications of this nature debutanized, light, straight-run gasoline is deheptanized, either in existing equipment, such as a catalytic reformer feed-preparation unit, or in new equipment. The C -C fraction is then split, and the C s, including debutanized C reactor efliuent, are processed to produce an isopentane product and a normal-pentane reactor feed. The degree of fractionation determines the product octane number, since normal pentane is recycled to extinction.

The C fraction can be employed in gasoline blending, or can be isomerized by one of two methods. Hexane fractions high in normal hexane content can be improved considerably by direct single-pass isomerization. Further improvement in octane number is possible by first splitting the isofrom the normal hexane, and then isomerizing the normal hexane fraction. Further octane improvement is possible by recycling normal hexane to extinction. This, however, would require an extra fractionation step to prevent an excessive build-up of methyl cyclopentane in the recycle stream.

An alternate method for processing normal pentane and the total hexane fraction in a single reactor involves deheptanizing la debutanized feed stock. The deheptanized feed is deisopentanized, and the resultant stream passed through the reaction system. The debutanized reactor efliuent then is fractionated to produce an isomerized hexane fraction and a pentane recycle stream which passes to the the deisopent-anizer. In this alternate processing, reaction condtions are determined by the more reactive hexanes, resulting in a lower conversion per pass of normal pentane. The greater fractionation cost must be balanced by the decreased reactor section costs, since only one reaction section is required.

The relative quantities of pentanes and hexanes, as well as the isoto normal-hexane ratio determine which processing method is most economical.

Results obtained using rhodium-promoted, silicaalumina catalysts show that no advantage in catalyst activity is gained by incorporating more than about 0.1% by weight of rhodium in the catalyst. However, in order to lengthen process cycles at high catalyst activity, larger amounts of rhodium, up to about 0.4% by weight of the catalyst, may be preferred when feedstocks containing considerable amounts of sulfur, or other potential poisons, are processed. Such catalysts are economical as compared with other noble-metal catalysts because of the low metal content required to obtain high maximum activity. The activity can be further increased by incorporating in the silica-alumina support, or catalyst, about 1-3% by weight of fluorine. This can be readily accomplished by incorporating in the catalyst support aqueous hydrofluoric acid, trifluoracetic or other fluoroorganic acid and drying the catalyst support.

It will be apparent to those skilled in the art that these modifications can be made without departing from the scope of this invention as defined by the appended claims.

We claim as our invention:

1. A hydroisomerization process which comprises processing substantially pure dry hydrogen and charging stock consisting essentially of saturated hydrocarbons within the range of 4-7 carbon atoms per molecule, said hydrocarbons being predominantly normal paraffinic, at a temperature within the range of about 600- 775 F., at isomerizing conditions of pressure and H hydrocarbon mol ratio in the presence of an isomerization catalyst consisting essentially of a silica-alumina, hydrocarbon-cracking catalyst having incorporated therein about 0.1% w. of rhodium, said silica-alumina catalyst containing not less than 50% by weight of silica.

2. Process in accordance with claim 1 in which the silica-alumina, hydrocarbon-cracking catalyst contains about 5095% by weight of silica and 505% by weight of alumina, based on said silica-alumina catalyst.

3. A hydroisomerization process which comprises processing substantially pure dry hydrogen and charging stock consisting essentially of saturated hydrocarbons within the range of 56 carbon atoms per molecule, said hydrocarbons being predominantly normal parafiinic, at a temperature within the range of about 600-775 F., a pressure within the range of about 1004000 p.s.i.g., and a hydrogen/hydrocarbon mol ratio within the range of about 0.5-5 in the presence of a catalyst consisting essentially of a silica-alumina hydrocarboncracking catalyst containing not less than about 50% nor more than by weight of silica and having incorporated therein about 0.1% by weight, based on total catalyst composition, of rhodium.

4. Process in accordance with claim 3 in which said feed stock consists predominantly of normal pentane.

5. Process in accordance with claim 3 in which said feedstock consists predominantly of normal hexane.

6. Process in accordance with claim 3 in which said feed stock consists predominantly of a mixture of normal pentane and normal hexane.

References Cited in the file of this patent UNITED STATES PATENTS 2,698,829 Haensel Jan. 4, 1955 2,841,626 Holzman et al. July 1, 1958 FOREIGN PATENTS 487,392 Canada Oct. 21, 1952 

1. A HYDROISOMERIZATION PROCESS WHICH COMPRISES PROCESSING SUBSTANTIALLY PURE DRY HYDROGEN AND CHARGING STOCK CONSISTING ESSENTIALLY OF SATURATED HYDROCARBONS WITHIN THE RANGE OF 4-7 CARBON ATOMS PER MOLECULE, SAID HYDROCARBONS BEING PREDOMINANTLY NORMAL PARAFFINIC, AT A TEMPERATURE WITHIN THE RANGE OF ABOUT 600775*F., AT ISOMERIZING CONDITIONS OF PRESSURE AND H2 HYDROCARBON MOL RATIO IN THE PRESENCE OF AN ISOMERIZATION CATALYST CONSISTING ESSENTIALLY OF A SILICA-ALUMINA, HYDROCARBON-CRACKING CATALYST HAVING INCORPORATED THEREIN ABOUT 0.1% W. OF RHODIUM, SAID SILICA-ALUMINA CATALYST CONTAINING NOT LESS THEN 50% BY WEIGHT OF SILICA. 